Methods And Apparatus For Producing Aromatics From Coal

ABSTRACT

A method for converting coal into BTX in which feed coal is converted to a 600-700° F.− product stream by direct liquefaction. This product stream is hydrocracked and hydroprocessed to produce a 350° F.− stream which in turn is fractionated to produce a 160° F.− stream and a 160/350° F. stream that contains 85-90% naphthenes. The 160/350° F. stream is catalytically reformed to produce an aromatic stream and a 160° F.− paraffinic stream. The aromatics stream can be separated into benzene toluene and xylene streams by distillation.

FIELD OF THE INVENTION

The present invention relates to direct coal liquefaction processes forefficiently producing high-value aromatics from coal.

BACKGROUND OF THE INVENTION

Current and presently proposed methods for the production of high valuearomatics (benzene, toluene, and xylene, i.e. BTX) from coal includeeither the capture of gaseous byproducts from the pyrolysis of coal inan airless environment, or the proposed conversion of coal to methanolfollowed by the conversion of the methanol to aromatics.

The pyrolysis method is of economic interest only as a byproduct ofanother process, e.g., producing coke from coal. The methanol toaromatics (MTA) process would involve the gasification of the coal feed,typically by partial oxidation (PDX), to produce syngas as the feed formethanol synthesis, converts coal to methanol and then converts methanolto aromatics by processing over a fixed or fluid bed of catalyst. Apublished estimate stated that the required investment for a 1 millionmetric ton per stream day MTA plant in China would be about $4.6billion. This high cost results in major part because of the need to useexpensive PDX for gasifying all of the coal. Available publicationsindicate that about 3 MT of coal would be required to produce 1 MT ofmethanol. A separate report states that 1.42 to 1.59 MT of methanolwould be required to produce 1 MT of aromatics. If these two numbers arecombined, it requires 4.26 to 4.77 MT of coal to produce 1 MT ofaromatics. It is estimated that the thermal efficiency of the MTAprocess is about 40 to 45%.

SUMMARY OF THE INVENTION

In accordance with the invention, a highly efficient and lower costmethod and system for producing high-value aromatics from coal isprovided in which the feed coal is converted by direct coal liquefaction(DCL) to a 1000° F.−, preferably an 800° F.−, more preferably a 600-750°F.− product, most preferably a 600-700° F.− product, at least the 350°F.+ portion of which is then hydrocracked to produce a 350° F.− productstream. The DCL and hydrocracked 350° F.− product streams are thenhydroprocessed to remove sulfur, nitrogen and oxygen compounds andfractionated into approximately 160° F.− and 160° F.+ output streams.The approximately 160/350° F. output stream is ideal for aromaticsproduction. It typically contains 85 to 90% naphthenes, 5 to 10%paraffins, plus some single ring aromatics, and the naphthenes areeasily converted into BTX in a catalytic reformer. The product from thecatalytic reformer can be processed in a solvent extraction unit toproduce a pure aromatics product and a paraffinic raffinate.Alternatively, after converting the naphthenes and heavier paraffins toaromatics in the catalytic reformer, the entire product can be passedover a bed of cracking catalyst in the same unit. This bed isomerizeslower value ethyl benzene to more valuable para-xylene and benzene, andcracks the remaining paraffins into an approximately 160° F.− productthat can be separated from the aromatics in the distillation tower,thereby producing a higher value aromatic product and eliminating theneed for solvent extraction.

The preferred DCL system includes a slurry DCL reactor containing amolybdenum or iron, preferably molybdenum, microcatalyst and is operatedat high conversion with the product boiling above the 600-700° F. rangepreferably being recycled and mixed with the DCL feed coal as anon-donor stream in a ratio of non-donor stream to coal at the input tothe reactor (on a moisture free weight basis) of between 1.6 and 3.5:1.By “non-donor” is meant that the recycle stream has not been processedin a hydrotreater to partially hydrogenate multi-ring aromatic compoundsin the stream in order to produce compounds that can donate hydrogenduring liquefaction.

In order to provide the additional hydrogen required for the DCL andhydrocracking processes, the bottoms from the DCL reactor can begasified in a PDX reactor. As an alternative, the additional hydrogenmay be provided by processing 160° F.− product of the upgrading and thecatalytic reformer by liquid PDX or steam naphtha reforming (SNR). Ifnatural gas is available, it can be used as the feed to the liquid PDXor SNR instead of the 160° F.− product.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 is a diagram of the flow stream of an embodiment of the methodand apparatus of the invention in which the additional hydrogen isprovided by PDX.

FIG. 2 is a schematic diagram of a preferred direct coal liquefactionsystem suitable for use in the illustrated embodiments of the invention.

FIG. 3 is a diagram of the flow stream of an embodiment of the methodand apparatus of the invention in which the additional hydrogen isproduced from the 160° F.− product stream.

FIG. 4 is a diagram of the flow stream of an embodiment of the methodand apparatus of the invention in which the additional hydrogen isproduced from natural gas feed.

DETAILED DESCRIPTION OF ILLUSTRATED EMBODIMENTS

Referring now to FIG. 1 of the drawings, there is illustrated aschematic of the overall flow scheme of a first embodiment of a coal toaromatics plant 100 according to the invention. The coal feed 101 issupplied to the DCL unit 103 that is preferably operated at a highconversion of 80+% on a moisture and ash free (MAF) basis. In the DCLunit 103, coal is hydrogenated to produce an approximately a 1000° F.−,preferably an 800° F.−, more preferably a 600-750° F.− product, mostpreferably a 600-700° F.− product stream 105 and an effluent stream 107that consists of ash, unconverted coal, and liquids boiling above 1000°F. The product stream 105 flows to the upgrader 109, in which at leastthe 350° F.+ portion of the stream is hydrocracked to produce a 350° F.−product stream. The 350° F.− product stream from the hydrocracker, or,if only the 350° F.+ portion of the DCL product was hydrocracked, thecombined 350° F.− DCL and hydrocracked streams are hydroprocessed toremove heteroatoms, and to convert most of the aromatics present intonaphthenes to produce a 350° F.− product stream 113. This stream is thenfractionated in the atmospheric fractionator 115 into approximately 160°F.− and 160° F.+ outputs 117 and 119, respectively. Light gases producedin the DCL unit 103 and the upgrader 109 may be used to supply a portionof the fuel for the plant. Excess light gases (C2−) may be sent to asteam methane reformer to supply a portion of the hydrogen required bythe DCL unit 103. The bottoms stream 107 from the DCL unit 103 isgasified in the PDX unit 111 for supplying additional hydrogen to theDCL unit 103 and the upgrader 109.

The approximately 160/350° F. stream 119 from the fractionator 115 istypically made up of 85% to 90% naphthenes, 5 to 10% paraffins, and somesingle ring aromatics. This stream is fed to the catalytic reformer 121where the naphthenes are converted into aromatics. The output of thecatalytic reformer 121 is fed to the solvent extractor unit 123 where itis separated into a pure aromatics stream 125 and a lower boiling pointparaffinic raffinate stream 127. The aromatics stream 125 can beseparated into its benzene, toluene and xylene components bydistillation.

As attractive alternative to the use of the extraction unit 123 is thatdescribed in U.S. Pat. No. 5,472,593 in which, after converting thenaphthenes and heavier paraffins to aromatics in the catalytic reformer,the entire product is passed over a bed of catalyst in the same unitcomprising a medium-pore molecular sieve having a pore size of from 5 to6.5 .ANG., a refractory inorganic oxide, a platinum-group metalcomponent and a lead or bismuth metal attenuator. This bed converts thelower value ethyl benzene to more valuable para-xylene and benzene, andcracks the remaining paraffins into a 160° F.− product that can beseparated from the aromatics by distillation, thereby producing a highervalue aromatic product and eliminating the need for solvent extraction.This alternative does, however, require additional hydrogen forcracking. The disclosure of U.S. Pat. No. 5,472,593 is herebyincorporated by reference in its entirety.

CO₂ produced by the PDX unit 111, and optionally, by the DCL reactorsystem 103 and/or other components of the liquefaction and upgradingsystem, is fed to the algae production system 129, which includes aphoto-bio reactor (PBR) in which the CO₂ is used to produce preferablyblue-green algae through photosynthesis. The DCL reactor system 103 andespecially the upgrading system 109 also produce NH₃, which can be fedto the algae production system 129 as a nutrient. The algae from thealgae production system 129 is preferably used to produce abiofertilizer 131. Methods and apparatus suitable for use in the presentinvention for producing algae and biofertilizer are disclosed in U.S.patent application Ser. No. 13/316,546 that was filed on Dec. 11, 2011,the disclosure of which is hereby Incorporated by reference in itsentirety.

Referring now to the embodiment of a DCL system illustrated in FIG. 2 ofthe drawings, the coal feed is dried and crushed in a conventional gasswept roller mill 201 to a moisture content of 1 to 4%. Crushed anddried coal is fed into a mixing tank 203 where it is mixed with a streamconstituted by a 600 to 700° F.+ fraction, preferably a 650° F.+fraction, of the output of the liquefaction reactor to form a slurrystream. The catalyst precursor in the illustrated embodiment preferablyis in the form of an aqueous water solution of phosphomolybdic acid(PMA) in an amount that is equivalent to adding between 50 wppm and 2%molybdenum relative to the dry coal feed. In the slurry mix tank 203,typical operating temperature ranges from 300 to 600° F. and morepreferably between 300 and 500° F. From the slurry mix tank, thecatalyst containing slurry is delivered to the slurry pump 205. Theselection of the appropriate mixing and temperature conditions is basedon experimental work quantifying the rheological properties of thespecific slurry blend being processed.

Most of the remaining moisture in the coal is driven off in the mixingtank due to the hot atmospheric fractionator bottoms feeding to themixing tanks. Residual moisture and any entrained volatiles arecondensed out as sour water (not shown in FIG. 2). The coal in theslurry leaving the mixing tank 203 has about 0.1 to 1.0% moisture. Theslurry formed by the coal, 600 to 700 to 1,000° F. stream from thevacuum fractionator 221, and the 600 to 700° F.+ stream fraction fromthe atmospheric fractionator 219 is pumped from the mixing tank 203 andthe pressure is raised to about 2,000 to 3,000 psig (138 to 206 kg/cm²g) by the slurry pumping system 205. The resulting high pressure slurrymay be preheated in a heat exchanger (not shown), mixed with a treat gasconsisting of recycled and makeup treat gas containing over 80%hydrogen, and then further heated in furnace 207.

The coal slurry and hydrogen mixture is fed to the input of the firststage of the series-connected liquefaction reactors 209, 211 and 213 atbetween 600 to 700° F. (316 to 371° C.) and 2,000 to 3,000 psig (138 to206 kg/cm² g). The reactors 209, 211 and 213 are simple up-flow tubularvessels, the total length of the three reactors being 40 to 200 feet.The temperature rises from one reactor stage to the next as a result ofthe highly exothermic coal liquefaction reactions. In order to maintainthe maximum temperature in each stage below about 800 to 900° F. (427 to482° C.), a portion of the hydrogen based treat gas is preferablyinjected between reactor stages. The hydrogen partial pressure in eachstage is preferably maintained at a minimum of about 1,000 to 2,000 psig(69 to 138 kg/cm² g).

The effluent from the last stage of liquefaction reactor is separatedinto a gas stream and a liquid/solid stream, and the liquid/solid streamlet down in pressure, in the separation and cooling system 215. The gasstream is cooled to condense out the liquid vapors of H2O, naphtha,distillate, and solvent. The remaining gas is then processed to removeH₂S, NH₃ and CO₂

Most of the processed gas is then sent to a hydrogen recovery system,not shown, for further processing by conventional means to recover thehydrogen contained therein, which is then recycled to be mixed with thecoal slurry. The remaining portion of the processed gas is purged toprevent buildup of light ends in the recycle loop. Hydrogen recoveredtherefrom can be used in the downstream hydro-processing upgradingsystem.

The depressurized liquid/solid stream and the hydrocarbons condensedduring the gas cooling are sent to the atmospheric fractionator 219where they are separated into light ends and in the preferredembodiment, a 600 to 700° F.− fraction, and a 600 to 700° F.+ fraction.The light ends are processed to recover hydrogen and C₁-C₂ hydrocarbonsthat can be used for fuel gas and other purposes. The 600 to 700° F.−fraction is sent to upgrading for aromatics production. Alternatively,the fractionator 219 could be arranged to produce 1000° F.− and 1000°F.+ fractions or 800° F.− and 800° F.+ fractions in which the 1000° F.−or 800° F.− fraction would be sent to the upgrading step.

In the preferred embodiment, a portion of the 600 to 700° F.+ (316 to371° C.+) is recycled to the slurry mix tank. The remaining 600 to 700°F.+ fraction produced from the atmospheric fractionator 219 is fed tothe vacuum fractionator 221 wherein it is separated into a 1000° F.−fraction and a 1000° F.+ fraction. The 1000° F.− fraction is added tothe 600 to 700° F.+ stream being recycled to the slurry mix tank 203.

In the preferred embodiment illustrated in FIG. 2, the 1000° F.+fraction from the vacuum fractionator 221 is sent to be gasified by thePDX system 223 (shown in FIG. 1 as PDX 111) to generate hydrogen for usein the liquefaction and upgrading. Alternatively, instead of the PDXsystem 223, the 1000° F.+ bottoms from the vacuum fractionator 221 maybe processed in a Circulating Fluid Bed boiler, a cement plant, or soldas a feed for asphalt paving or electrode manufacture. G. E., Shell, andothers offer commercial processes for gasification (partial oxidation)of the 1000° F.+ bottoms and Circulating Fluid Bed boiler manufacturessuch as Foster-Wheeler and Alstom offer technology for combusting the1000° F.+ bottoms.

Catalysts useful in DCL processes also include those disclosed in U.S.Pat. Nos. 4,077,867, 4,196,072 and 4,561,964, the disclosures of whichare hereby incorporated by reference in their entirety. Other DCLreactor systems suitable for use in the process of the invention aredisclosed in U.S. Pat. Nos. 4,485,008, 4,637,870, 5,200,063, 5,338,441,and 5,389,230, and U.S. patent application Ser. No. 13/657,087,thedisclosures of which are hereby incorporated by reference in theirentirety.

The preferred DCL Process combines several elements that contribute tomaximum BTX Product production and maximum thermal efficiency. Theseinclude, very importantly, the recycle of a non-donor 600 to 700° F.+stream, preferably including atmospheric fractionator bottoms, tomaintain a ratio of the recycle stream to coal at the input to thereactors 209, 211, 213 that is between 1.6:1 and 3.5:1 on a moisturefree weight basis; the use of a microcatalyst in the form of finelydivided molybdenum; and the use of a much lower treat gas rate than inprevious systems. Also, the use of bottoms recycle, and multiple slurryreactors in series contribute to the benefits of the process.

Use of a microcatalyst, which is either a compound of molybdenum oriron, more preferably molybdenum, and added at 100 to 1,000 wppm, morepreferably 100 to 500 wppm, and most preferably 100 to 300 wppm,eliminates several disadvantages to the use of a donor solvent such asrequired by prior DCL systems. First, energy is lost during preparationof the donor solvent. Second, energy is required to preheat the donorsolvent in the solvent hydrotreater and hydrogen must be compressed andcirculated around the hydrotreater. Thirdly, the heat release duringpartial hydrogenation of the donor solvent is lost during cooling priorto separation of hydrogen for recycle. In comparison, all of the heatrelease occurs in the liquefaction reactors during operation with a 600to 700° F.+ recycle stream, which minimizes the preheat requirementprior to liquefaction. These factors contribute to the higher thermalefficiency of the microcatalytic coal liquefaction process. Moreover,the use of a microcatalyst and the consequent elimination of the needfor a donor solvent also eliminates the need for an expensive solventhydrotreater to generate the donor solvent, thereby substantiallyreducing the capital and operating cost of the system. It also permitsthe use of coals having substantially higher ash contents, from 6 to 20wt % or more on a moisture free basis, and the recycle of asubstantially higher portion of bottoms than were possible with donorsolvent systems. Examples of microcatalysts and their method ofpreparation are described in U.S. Pat. No. 4,226,742, the contents ofwhich are hereby incorporated by reference in their entirety.

The 600 to 700° F.+ fraction recycled from the atmospheric fractionator219 and the 1000° F.− fraction from the vacuum fractionator 221 as thenon-donor stream being recycled to the slurry mix tank 203 providespreheat for the coal and solvent in the slurry mix tank 203. This raisesthe temperature in the mix tank to 300° F. to 500° F., more preferably350° F. to 500° F., and most preferably about 400 to 500° F. Thisfurther reduces the energy requirement for preheating the slurry priorto liquefaction. A significant portion of the of the microcatalyst isentrained in the 600 to 700° F.+ fraction recycled from the atmospherictower 219, so that recycling a larger portion of such fraction increasesthe catalyst concentration in the DCL reactors 209, 211, 213, therebydecreasing the requirement for the addition of fresh catalyst precursorand increasing the conversion efficiency of the DCL process.

Use of the non-donor 600° F. to 700° F.+ stream, more preferably 630° F.to 670° F+, and most preferably a 650° F+, process derived recyclesolvent in the DCL process reduces cracking, relative to donor solvent,and produces a 650° F.− product with a greater fraction of diesel andless light gases and naphtha. The 650° F.− product can be selectivelyupgraded to finished products in fixed bed upgrading reactors.

The much lower treat gas rate of 600 to 900 NL per kg of slurry has asignificant impact on thermal efficiency, plant investment, andoperating cost. The required recycle treat gas rate for the DCL processof the invention is up to three times lower than the preferred gas ratein the NEDOL program (without taking into account the treat gas rate tothe solvent hydrotreater, which makes the difference even larger). Thishas an important impact on power requirements for the compressor andfuel requirements for slurry preheat furnace 207 and solventhydrotreater preheat.

The use of two to four, more preferably three slurry reactors in seriesapproaches a plug flow reactor and hence has as little as two thirds ofthe required volume of one or two ebullated bed reactors such as used insome prior DCL systems. Since all of the heat is released in the threeliquefaction reactors, the temperature profile can be also maintained tomaximize selectivity to liquids. Operation of the initial reactor at asomewhat lower temperature has been reported in previous patents as aroute to increase conversion and liquid yields.

An exemplary process for upgrading the liquid product of the DCLreactors 209, 211, 213 is disclosed in U.S. Pat. No. 5,198,099, thedisclosure of which is hereby incorporated by reference in its entirety.Other processes and systems suitable for upgrading the liquid productsare commercially available from vendors such as UOP, Axens, Criterionand others.

Referring now to FIG. 3 of the drawings, there is illustrated a secondembodiment of the coal to aromatics flow scheme of the invention.Elements of the flow scheme that are the same as corresponding elementsof the embodiment of FIG. 1 are identified with the same referencenumbers as the corresponding elements of FIG. 1. The primary differencebetween the embodiments of FIGS. 1 and 3 is that, in FIG. 3, theadditional hydrogen required for the DCL system 103 and the upgrader 109is produced by the H₂ Plant 303 rather than by the PDX unit 111 used inFIG. 1. The H₂ Plant 303 can be implemented as a steam naptha reformer(SNR) or a liquid PDX unit, both of which are well known standardequipment in the art. For instance, SNR's are available from sourcessuch as Akzo Nobel N.V. Similarly, liquid PDX units are available fromsources such as Haldor-Topsoe, Inc. or Lurgi GmbH. In either case, atleast the C4− portion and if needed, part of the lighter portion of theC5+ portion of the 160° F.− stream from the solvent extractor 123 isused as the feed to the H₂ Plant 303.

Referring now to FIG. 4 of the drawings, there is illustrated a thirdembodiment of the coal to aromatics process and system of the inventionin which components that are the same as corresponding components inFIGS. 1 or 3 are labeled with the same reference numbers as suchcorresponding components in FIGS. 1 or 3. The primary difference betweenthe embodiments of FIGS. 4 and 3 is that, in FIG. 4, the feed to the H₂Plant 303 that produces the additional hydrogen required for the DCLsystem 103 and the upgrader 109 is supplied by natural gas rather thanby a the C1-C4 portion of the output from the solvent extractor 123. Inthis case the H₂ Plant 303 is implemented as an SMR. The feed to the H₂Plant 303 can also be from sources such as shale gas, or coal minemethane. The SMR technology is utilized worldwide in refineries and isoffered by many commercial vendors such as Haldor-Topsoe.

The embodiment of the invention illustrated in FIG. 1 is less expensiveto implement and has a substantially higher thermal efficiency (between60 and 65%) than the prior MTA systems (40-45%). In the configuration ofFIG. 1, PDX (of the DCL bottoms) is still required, but the quantity ofmaterial being processed in the PDX unit is greatly reduced. Since theprinciple product is aromatics, the remaining exported byproduct is aC3/160° F. stream from the solvent extractor 123. This stream would beconverted to lighter paraffins if the alternative approach disclosed inU.S. Pat. No. 5,472,593 is utilized, but will also require additionalhydrogen for cracking.

The embodiment of the invention illustrated in FIG. 3 requires thelowest investment per ton of aromatics and also has the simplest flowscheme. Instead of bottoms PDX, the lighter portion of the low valueC1/160° F. product is utilized as a feed for the production of H₂. Thisallows use of lower cost H₂ generating technologies including liquid PDX(no ash) or SNR. A plant implementing the flow scheme of FIG. 3 consumesmuch of the lower value products that it produces, and therefore has alower thermal efficiency (about 51.5%) than the embodiment of FIG. 1.Thus, there is a trade-off between higher thermal efficiency (FIG. 1),and lower investment, simpler, easier to operate plant (FIG. 3).

The embodiment of FIG. 4 is preferable where inexpensive natural gas isavailable. In this embodiment, H₂ is generated via SMR and the DCLbottoms are preferably sent to a CFB power plant for generation ofpower. Given the lower cost of shale or natural gas, this embodimentwould be economically preferable to the production of H₂ from the 160°F.− stream. This stream can better be sent to a Steam Cracking systemfor the production of olefins or aromatics. This embodiment also has asubstantially higher thermal efficiency (about 70%).

1. A method for producing aromatics from coal comprising the steps of:a. producing a 1000 F− product stream from feed coal by direct coalliquefaction (DCL); b. hydrocracking the 1000° F.− product stream toproduce a 350° F.− product stream; c. fractionating the 350° F.− productstream into 160/350° F. and lower boiling point streams; and d.converting the 160/350° F. stream to aromatics and a lower boiling pointstream by catalytic reforming.
 2. The method of claim 1 wherein thefraction produced in step a is an 800° F.− fraction.
 3. The method ofclaim 1 wherein the fraction produced in step a is a 600-700° F.−fraction.
 4. The method of claim 1 further including gasifying bottomsproduced in the direct coal liquefaction to produce hydrogen for supplyto the direct coal liquefaction and hydrocracking steps.
 5. The methodof claim 1 or 3 further including processing at least a portion of thelower boiling point streams to produce hydrogen for supply to the directcoal liquefaction and/or hydrocracking steps.
 6. The method of claim 1further including processing natural gas feed to produce hydrogen forsupply to the direct coal liquefaction and hydrocracking steps.
 7. Themethod of claim 1 further including separating aromatics produced instep d. into benzene toluene and xylene streams.
 8. The method of claim1 or 3 wherein the lower boiling point stream in step c is a 160° F.−stream.